Two-stage hydrocracking process for producing naphtha, comprising a hydrogenation stage implemented downstream of the second hydrocracking stage

ABSTRACT

The present invention is based on the use of a two-step hydrocracking process for the production of naphtha, comprising a step of hydrogenation placed downstream of the second hydrocracking step, the hydrogenation step treating the effluent resulting from the second hydrocracking step in the presence of a specific hydrogenation catalyst. Furthermore, the hydrogenation step and the second hydrocracking step are performed under specific operating conditions and in particular under quite specific temperature conditions.

TECHNICAL FIELD

The invention relates to a two-step hydrocracking process for removingheavy polycyclic aromatic compounds (HPNAs) without reducing the yieldof upgradable products.

Hydrocracking processes are commonly used in refinery for transforminghydrocarbon mixtures into readily upgradable products. These processesmay be used to transform light cuts, for instance petroleums, intolighter cuts (LPG). However they are customarily used more forconverting heavier feedstocks (such as heavy synthetic or petroleumcuts, for example gas oils resulting from the vacuum distillation oreffluents from a Fischer-Tropsch unit) into petroleum or naphtha,kerosene, gas oil.

Certain hydrocracking processes make it possible to also obtain a highlypurified residue that may constitute excellent bases for oils. One ofthe effluents that is particularly targeted by the hydrocracking processis middle distillate (fraction which contains the gas oil cut and thekerosene cut), i.e. cuts with an initial boiling point of at least 150°C. and with a final boiling point below the initial boiling point of theresidue, for example below 340° C., or else below 370° C. The LightPetroleum or Light Naphtha cut (having an initial boiling point above20° C. and a final boiling point below 80° C.), and the Heavy Petroleumor Heavy Naphtha cut (having an initial boiling point above 70° C. and afinal boiling point below 250° C.), are also desired for uses in fuelbases or for petrochemistry, and certain hydrocracking processes aredesigned to maximize the production of the “Heavy Naphtha” cut.

Hydrocracking is a process which draws its flexibility from three mainelements, namely: the operating conditions used, the types of catalystsemployed, and the fact that the hydrocracking of hydrocarbon feedstocksmay be performed in one or two steps.

In particular, the hydrocracking of vacuum distillates or VDs makes itpossible to produce light cuts (gas oil, kerosene, naphthas, and thelike) which are more upgradeable than the VD itself. This catalyticprocess does not make it possible to completely convert the VD intolight cuts. After fractionation, there thus remains a more or lesssignificant proportion of unconverted VD fraction, referred to as UCO orUnConverted Oil. To increase the conversion, this unconverted fractionmay be recycled to the inlet of the hydrotreating reactor or to theinlet of the hydrocracking reactor in the case of a one-stephydrocracking process or to the inlet of a second hydrocracking reactortreating the unconverted fraction at the end of the fractionating step,in the case of a two-step hydrocracking process.

It is known that the recycling of said unconverted fraction resultingfrom the fractionating step to the second hydrocracking step of atwo-step process results in the formation of heavy (polycyclic) aromaticcompounds referred to as HPNAs during the cracking reactions and thus inthe undesirable accumulation of said compounds in the recycle loop,resulting in the degradation of the performance of the catalyst of thesecond hydrocracking step and/or in the fouling thereof. A purge isgenerally installed on the recycle loop of said unconverted fraction, ingeneral on the line at the bottom of the fractionation, in order toreduce the concentration, in the recycle loop, of HPNA compounds, thepurge flow rate being adjusted so as to balance the formation flow ratethereof. Specifically, the heavier the HPNAs, the greater their tendencyto remain in this loop, to accumulate, and to grow heavier.

However, the overall conversion of a two-step hydrocracking process isdirectly linked to the amount of heavy products purged at the same timeas the HPNAs. This purge therefore leads to a loss of upgradableproducts which are also extracted with the HPNAs via this purge.

Depending on the operating conditions of the process, said purge may bebetween 0 and 5% by weight of the unconverted heavy fraction (UCO)relative to the incoming VD mother feedstock, and preferably between0.5% and 3% by weight. The yield of upgradable products is thereforereduced accordingly, which constitutes a not inconsiderable economicloss for the refiner.

Throughout the remainder of the text, the HPNA compounds are defined aspolycyclic or polynuclear aromatic compounds which therefore compriseseveral fused benzene nuclei or rings. They are customarily referred toas PNAs, Polynuclear Aromatics, for the lightest of them and as HPAs orHPNAs, Heavy PolyNuclear Aromatics, for the compounds comprising atleast seven aromatic nuclei (for instance coronene, composed of 7aromatic rings). These compounds, formed during undesirable secondaryreactions, are stable and very difficult to hydrocrack.

PRIOR ART

There are various patents that relate to processes which seek tospecifically treat the problem linked to HPNAs so that they are notdetrimental to the process simultaneously in terms of performance, cycletime and operability.

Certain patents claim the elimination of HPNA compounds byfractionation, distillation, solvent extraction or adsorption on atrapping mass (WO2016/102302, U.S. Pat. Nos. 8,852,404 9,580,663,5,464,526 and 4,775,460).

Another technique consists in hydrogenating effluents containing theHPNAs in order to limit the formation and accumulation thereof in therecycle loop.

U.S. Pat. No. 3,929,618 describes a process for hydrogenating andopening the rings of hydrocarbon feedstocks containing fused polycyclichydrocarbons in the presence of a catalyst based on NaY zeolite andexchanged with nickel.

U.S. Pat. No. 4,931,165 describes a one-step hydrocracking process withrecycling comprising a step of hydrogenation over the recycle loop ofthe gases.

U.S. Pat. No. 4,618,412 describes a one-step hydrocracking process inwhich the unconverted effluent resulting from the hydrocracking stepcontaining HPNAs is sent to a step of hydrogenation over a catalystbased on iron and on alkali or alkaline-earth metals, at temperatures ofbetween 225° C. and 430° C. before being recycled into the hydrocrackingstep.

U.S. Pat. No. 5,007,998 describes a one-step hydrocracking process inwhich the unconverted effluent resulting from the hydrocracking stepcontaining HPNAs is sent to a step of hydrogenation over a zeolitichydrogenation catalyst (zeolite with pore sizes between 8 and 15 Å) alsocomprising a hydrogenation component and a clay.

U.S. Pat. No. 5,139,644 describes a process similar to that of U.S. Pat.No. 5,007,998 with coupling to a step of adsorption of the HPNAs on anadsorbent.

U.S. Pat. No. 5,364,514 describes a conversion process comprising afirst hydrocracking step, the effluent resulting from this first stepthen being split into two effluents. A portion of the effluent resultingfrom the first hydrocracking step is sent to a second hydrocracking stepwhile the other portion of the effluent resulting from the firsthydrocracking step is sent simultaneously to a step of hydrogenation ofaromatics using a catalyst comprising at least one noble metal fromgroup VIII on an amorphous or crystalline support. The effluentsproduced in said hydrogenation step and second hydrocracking step arethen sent to the same separation step or to dedicated separation steps.

Patent application US2017/362516 describes a two-step hydrocrackingprocess comprising a first hydrocracking step followed by fractionationof the hydrocracked stream producing an unconverted effluent comprisingHPNAs which is recycled and referred to as the recycle stream. Thisrecycle stream is then sent to a hydrotreating step which enables thesaturation, by hydrogenation, of the HPNA aromatic compounds. Thishydrotreating step produces a hydrogenated stream which is then sent toa second hydrocracking step.

The essential criterion of the invention of US2017/362516 lies in thefact that the hydrotreating step that enables the hydrogenation of theHPNAs is located upstream of the second hydrocracking step. Thehydrotreating step and the second hydrocracking step may be performed intwo different reactors or in the same reactor. When they are performedin the same reactor, said reactor comprises a first catalytic bedcomprising a hydrotreating catalyst that enables the saturation of thearomatics, followed by catalytic beds comprising the hydrocrackingcatalyst of the second step.

The hydrotreating catalyst used is a catalyst comprising at least onegroup VIII metal and preferably a group VIII noble metal comprisingrhenium, ruthenium, rhodium, palladium, silver, osmium, iridium,platinum and/or gold, it being possible for said catalyst to optionallyalso comprise at least one non-noble metal and preferably cobalt,nickel, vanadium, molybdenum and/or tungsten, supported preferably onalumina. Other zeolitic catalysts and/or hydrogenation catalysts thatare not supported may be used.

The research studies performed by the Applicant have led the Applicantto discover an improved implementation of the hydrocracking processwhich makes it possible to limit the formation of HPNA in the secondstep of a two-step hydrocracking scheme and therefore to increase thecycle time of the process by limiting the deactivation of thehydrocracking catalyst. Another advantage of the present invention makesit possible to minimize the purge and therefore to maximize theupgradable products and in particular the yields of naphtha.

The present invention is based on the use of a two-step hydrocrackingprocess for the production of naphtha, comprising a step ofhydrogenation placed downstream of the second hydrocracking step, thehydrogenation step treating the effluent resulting from the secondhydrocracking step in the presence of a specific hydrogenation catalyst.Furthermore, the hydrogenation step and the second hydrocracking stepare performed under specific operating conditions and in particularunder quite specific temperature conditions.

SUMMARY OF THE INVENTION

In particular, the present invention relates to a process for producingnaphtha and in particular “heavy naphtha” from hydrocarbon feedstockscontaining at least 20% by volume and preferably at least 80% by volumeof compounds boiling above 340° C., said process comprising andpreferably consisting of at least the following steps:

a) a step of hydrotreating said feedstocks in the presence of hydrogenand at least one hydrotreating catalyst, at a temperature of between200° C. and 450° C., under a pressure of between 2 and 25 MPa, at aspace velocity of between 0.1 and 6 h⁻¹ and with an amount of hydrogenintroduced such that the litre of hydrogen/litre of hydrocarbon volumeratio is between 100 and 2000 Nl/l,

b) a step of hydrocracking at least one portion of the effluentresulting from step a), the hydrocracking step b) taking place, in thepresence of hydrogen and at least one hydrocracking catalyst, at atemperature of between 250° C. and 480° C., under a pressure of between2 and 25 MPa, at a space velocity of between 0.1 and 6 h⁻¹ and with anamount of hydrogen introduced such that the litre of hydrogen/litre ofhydrocarbon volume ratio is between 80 and 2000 Nl/l,

c) a step of high-pressure separation of the effluent resulting from thehydrocracking step b) to produce at least a gaseous effluent and aliquid hydrocarbon effluent,

d) a step of distilling at least one portion of the liquid hydrocarboneffluent resulting from step c) performed in at least one distillationcolumn, from which step the following are drawn off:

-   -   a gaseous fraction,    -   at least one fraction comprising the converted hydrocarbon        products having at least 80% by volume of products boiling at a        temperature below 250° C., preferably below 220° C., preferably        below 190° C. and more preferably below 175° C., and    -   an unconverted liquid fraction having at least 80% by volume of        products having a boiling point above 175° C., preferably above        190° C., preferably above 220° C. and more preferably above 250°        C.,

e) optionally a purging of at least one portion of said unconvertedliquid fraction containing HPNAs, having at least 80% by volume ofproducts having a boiling point above 175° C., before the introductionthereof into step f),

f) a second step of hydrocracking at least one portion of theunconverted liquid fraction having at least 80% by volume of productswith a boiling point above 175° C., resulting from step d) andoptionally purged, said step f) being performed in the presence ofhydrogen and of at least one second hydrocracking catalyst, at atemperature TR1 of between 250° C. and 480° C., under a pressure ofbetween 2 and 25 MPa, at a space velocity of between 0.1 and 6 h⁻¹ andwith an amount of hydrogen introduced such that the litre ofhydrogen/litre of hydrocarbon volume ratio is between 80 and 2000 Nl/l,

g) a step of hydrogenating of at least one portion of the effluentresulting from step f) taking place in the presence of hydrogen and ahydrogenation catalyst, at a temperature TR2 between 150° C. and 470°C., under a pressure of between 2 and 25 MPa, at a space velocity ofbetween 0.1 and 50 h⁻¹ and with an amount of hydrogen introduced suchthat the litre of hydrogen/litre of hydrocarbon volume ratio is between100 and 4000 Nl/l, said hydrogenation catalyst comprising at least onemetal from group VIII chosen from nickel, cobalt, iron, palladium,platinum, rhodium, ruthenium, osmium and iridium alone or as a mixtureand not containing any metal from group VIB and a support chosen fromrefractory oxide supports, and in which the temperature TR2 is at least10° C. below the temperature TR1,

h) a step of high-pressure separation of the effluent resulting from thehydrogenation step g) to produce at least a gaseous effluent and aliquid hydrocarbon effluent,

i) recycling, to said distillation step d), at least one portion of theliquid hydrocarbon effluent resulting from step h).

The temperature expressed for each step is preferably a weighted averagetemperature over all of the catalytic beds, or WABT, for example asdefined in the book “Hydroprocessing of Heavy Oils and Residua”, JorgeAncheyta, James G. Speight—2007—Science.

One advantage of the present invention is that it provides a two-stepprocess for hydrocracking a VD feedstock that makes it possiblesimultaneously to maximize the overall yield of said process in terms of“heavy naphtha” cut and to increase the cycle time of the process bylimiting the deactivation of the hydrocracking catalyst. The purge mayalso be minimized, which maximises the overall conversion of theprocess.

Throughout the remainder of the text, the term “heavy naphtha” fractionrefers to the heavy petroleum fraction resulting from the atmosphericdistillation at the outlet of the hydrocracker. Said factionadvantageously comprises at least 80% by volume of products boiling at aboiling point of between 70° C. and 250° C. and preferably between 75°C. and 220° C., preferably between 80° C. and 190° C. and morepreferably between 80° C. and 175° C.

The term “light naphtha” fraction refers to the light petroleum fractionresulting from the atmospheric distillation at the outlet of thehydrocracker. Said faction advantageously comprises at least 80% byvolume of products boiling at a boiling point of between 20° C. and 80°C., preferably between 25° C. and 75° C. and preferably between 30° C.and 70° C.

Feedstocks

The present invention relates to a process for hydrocracking hydrocarbonfeedstocks referred to as mother feedstock, containing at least 20% byvolume, and preferably at least 80% by volume, of compounds boilingabove 340° C., preferably above 350° C. and preferably between 350° C.and 580° C. (i.e. corresponding to compounds containing at least 15 to20 carbon atoms).

Said hydrocarbon feedstocks may advantageously be chosen from VGOs(vacuum gas oils) or vacuum distillates (VDs) or gas oils, for instancegas oils resulting from the direct distillation of crude or fromconversion units, such as FCC units (for example LCO or Light CycleOil), coker or visbreaking units, and also feedstocks originating fromunits for the extraction of aromatics from lubricating oil bases orresulting from the solvent dewaxing of lubricating oil bases, or elsedistillates originating from the desulfurization or hydroconversion ofATRs (atmospheric residues) and/or VRs (vacuum residues), or else thefeedstock may advantageously be a deasphalted oil, or feedstocksresulting from biomass or any mixture of the feedstocks mentionedpreviously, and preferably VGOs.

Paraffins resulting from the Fischer-Tropsch process are excluded.

The nitrogen content of the mother feedstocks treated in the processaccording to the invention is usually greater than 500 ppm by weight,preferably between 500 and 10 000 ppm by weight, more preferably between700 and 4000 ppm by weight and more preferably still between 1000 and4000 ppm by weight. The sulfur content of the mother feedstocks treatedin the process according to the invention is usually between 0.01% and5% by weight, preferably between 0.2% and 4% by weight and morepreferably still between 0.5% and 3% by weight.

The feedstock may optionally contain metals. The cumulative content ofnickel and vanadium of the feedstocks treated in the process accordingto the invention is preferably less than 1 ppm by weight.

The feedstock may optionally contain asphaltenes. The asphaltene contentis generally less than 3000 ppm by weight, preferably less than 1000 ppmby weight and even more preferably less than 200 ppm by weight.

In the case where the feedstock contains compounds of resin and/orasphaltene type, it is advantageous to pass the feedstock beforehandover a bed of catalyst or of adsorbent different from the hydrocrackingor hydrotreating catalyst.

Step a)

In accordance with the invention, the process comprises a step a) ofhydrotreating said feedstocks in the presence of hydrogen and at leastone hydrotreating catalyst, at a temperature of between 200° C. and 450°C., under a pressure of between 2 and 25 MPa, at a space velocity ofbetween 0.1 and 6 h⁻¹ and with an amount of hydrogen introduced suchthat the litre of hydrogen/litre of hydrocarbon volume ratio is between100 and 2000 Nl/l.

The operating conditions such as temperature, pressure, degree ofhydrogen recycling or hourly space velocity, may be highly variabledepending on the nature of the feedstock, on the quality of the productsdesired and on the plants which the refiner has at his disposal.

Preferably, the hydrotreating step a) according to the invention takesplace at a temperature of between 250° C. and 450° C., very preferablybetween 300° C. and 430° C., under a pressure of between 5 and 20 MPa,at a space velocity of between 0.2 and 5 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is between 300 and 1500 Nl/l.

Conventional hydrotreating catalysts may advantageously be used,preferably which contain at least one amorphous support and at least onehydro-dehydrogenating element chosen from at least one non-noble elementfrom Groups VIB and VIII, and usually at least one element from GroupVIB and at least one non-noble element from Group VIII.

Preferably, the amorphous support is alumina or silica/alumina.

Preferred catalysts are chosen from the catalysts NiMo, NiW or CoMo onalumina, and NiMo or NiW on silica/alumina.

The effluent resulting from the hydrotreating step and a portion ofwhich enters the hydrocracking step b) generally comprises a nitrogencontent preferably of less than 300 ppm by weight and preferably of lessthan 50 ppm by weight.

Step b)

In accordance with the invention, the process comprises a step b) ofhydrocracking at least one portion of the effluent resulting from stepa), and preferably all thereof, said step b) taking place, in thepresence of hydrogen and at least one hydrocracking catalyst, at atemperature of between 250° C. and 480° C., under a pressure of between2 and 25 MPa, at a space velocity of between 0.1 and 6 h⁻¹ and with anamount of hydrogen introduced such that the litre of hydrogen/litre ofhydrocarbon volume ratio is between 80 and 2000 Nl/l,

Preferably, the hydrocracking step b) according to the invention takesplace at a temperature of between 320° C. and 450° C., very preferablybetween 330° C. and 435° C., under a pressure of between 3 and 20 MPa,at a space velocity of between 0.2 and 4 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is between 200 and 2000 Nl/l.

In an embodiment that makes it possible to maximize the production of“heavy naphtha”, the operating conditions used in the process accordingto the invention generally make it possible to obtain conversions perpass, into products having at least 80% by volume of products havingboiling points below 250° C., preferably below 220° C., preferably below190° C. and more preferably below 175° C., of greater than 15% by weightand even more preferably of between 20% and 95% by weight.

The hydrocracking step b) according to the invention covers the pressureand conversion ranges extending from mild hydrocracking to high-pressurehydrocracking. The term “mild hydrocracking” refers to hydrocrackingwhich results in moderate conversions, generally of less than 40%, andwhich operates at low pressure, preferably between 2 MPa and 6 MPa.High-pressure hydrocracking is generally performed at greater pressures,between 5 MPa and 25 MPa, so as to obtain conversions of greater than50%.

The hydrotreating step a) and the hydrocracking step b) mayadvantageously be performed in the same reactor or in differentreactors. When they are performed in the same reactor, the reactorcomprises several catalytic beds, the first catalytic beds comprisingthe hydrotreating catalyst(s) and the following catalytic bedscomprising the hydrocracking catalyst(s).

Catalyst for the hydrocracking step b)

In accordance with the invention, the hydrocracking step b) is performedin the presence of at least one hydrocracking catalyst.

The hydrocracking catalyst(s) used in the hydrocracking step b) areconventional hydrocracking catalysts known to those skilled in the art,of bifunctional type combining an acid function with ahydrogenating-dehydrogenating function and optionally at least onebinder matrix. The acid function is provided by supports having a largesurface area (150 to 800 m²·g⁻¹ generally) having surface acidity, suchas halogenated (in particular chlorinated or fluorinated) aluminas,combinations of boron and aluminium oxides, amorphous silicas/aluminasand zeolites. The hydrogenating-dehydrogenating function is provided byat least one metal from group VIB of the Periodic Table and/or at leastone metal from group VIII.

Preferably, the hydrocracking catalyst(s) used in step b) comprise ahydrogenating-dehydrogenating function comprising at least one metalfrom group VIII chosen from iron, cobalt, nickel, ruthenium, rhodium,palladium and platinum, and preferably from cobalt and nickel.Preferably, said catalyst(s) also comprise(s) at least one metal fromgroup VIB chosen from chromium, molybdenum and tungsten, alone or as amixture, and preferably from molybdenum and tungsten.Hydrogenating-dehydrogenating functions of NiMo, NiMoW, NiW type arepreferred.

Preferably, the content of metal from group VIII in the hydrocrackingcatalyst(s) is advantageously between 0.5% and 15% by weight andpreferably between 1% and 10% by weight, the percentages being expressedas weight percentage of oxides relative to the total mass of catalyst.

Preferably, the content of metal from group VIB in the hydrocrackingcatalyst(s) is advantageously between 5% and 35% by weight andpreferably between 10% and 30% by weight, the percentages beingexpressed as weight percentage of oxides relative to the total mass ofcatalyst.

The hydrocracking catalyst(s) used in step b) may also optionallycomprise at least one promoter element deposited on the catalyst andchosen from the group formed by phosphorus, boron and silicon,optionally at least one element from group VIIA (chlorine and fluorinepreferred), optionally at least one element from group VIIB (manganesepreferred), and optionally at least one element from group VB (niobiumpreferred).

Preferably, the hydrocracking catalyst(s) used in step b) comprise atleast one amorphous or poorly crystallized porous mineral matrix ofoxide type chosen from aluminas, silicas, silica-aluminas, aluminates,alumina-boron oxide, magnesia, silica-magnesia, zirconia, titanium oxideor clay, alone or as a mixture, and preferably alumina orsilica-aluminas, alone or as a mixture.

Preferably, the silica-alumina contains more than 50% by weight ofalumina, preferably more than 60% by weight of alumina.

Preferably, the hydrocracking catalyst(s) used in step b) alsooptionally comprise a zeolite chosen from Y zeolites, preferably fromUSY zeolites, alone or in combination with other zeolites from amongbeta, ZSM-12, IZM-2, ZSM-22, ZSM-23, SAPO-11, ZSM-48 or ZBM-30 zeolites,alone or as a mixture. Preferably, the zeolite is USY zeolite alone.

When said catalyst comprises a zeolite, the content of zeolite in thehydrocracking catalyst(s) is advantageously between 0.1% and 80% byweight, preferably between 3% and 70% by weight, the percentages beingexpressed as percentage of zeolite relative to the total mass ofcatalyst.

A preferred catalyst comprises, and preferably consists of, at least onemetal from group VIB and optionally at least one non-noble metal fromgroup VIII, at least one promoter element, and preferably phosphorus, atleast one Y zeolite and at least one alumina binder.

An even more preferred catalyst comprises, and preferably consists of,nickel, molybdenum, phosphorus, a USY zeolite, and optionally also abeta zeolite, and alumina.

Another preferred catalyst comprises, and preferably consists of,nickel, tungsten, alumina and silica-alumina.

Another preferred catalyst comprises, and preferably consists of,nickel, tungsten, a USY zeolite, alumina and silica-alumina.

Step c)

In accordance with the invention, the process comprises a high-pressureseparation step c) comprising a separation means, for instance a seriesof disengagers at high pressure operating between 2 and 25 MPa, thepurpose of which is to produce a stream of hydrogen which is recycled bymeans of a compressor to at least one of the steps a), b), f) and/or g),and a hydrocarbon effluent produced in the hydrocracking step b) whichis preferentially sent to a steam stripping step preferably operating ata pressure of between 0.5 and 2 MPa, the purpose of which is to performseparation of the hydrogen sulfide (H₂S) dissolved in at least saidhydrocarbon effluent produced in step b).

Step c) allows the production of a liquid hydrocarbon effluent which isthen sent to the distillation step d).

Step d)

In accordance with the invention, the process comprises a step d) ofdistilling the effluent resulting from step c) to give at least agaseous fraction comprising the C1-04 light gases, a fraction comprisingthe converted hydrocarbon products having at least 80% by volume,preferably at least 95% by volume, of products boiling at a temperaturebelow 250° C., preferably below 220° C., preferably below 190° C. andmore preferably below 175° C., and an unconverted liquid fraction havingat least 80% by volume and preferably at least 95% by volume of productshaving a boiling point above 175° C., preferably above 190° C.,preferably above 220° C. and more preferably above 250° C.

Fractions having a boiling point that is between the boiling points ofthe “heavy naphtha” fraction and the unconverted heavy fraction may alsobe separated.

Optional Step e)

The process may optionally comprise a step e) of purging at least aportion of said unconverted liquid fraction containing HPNAs, resultingfrom the distillation step d).

Said purge is between 0 and 5% by weight of the unconverted liquidfraction relative to the feedstock entering said process, preferablybetween 0 and 3% by weight and very preferably between 0 and 2% byweight.

Step f)

In accordance with the invention, the process comprises a step f) ofhydrocracking said unconverted liquid fraction resulting from step d)and optionally purged in step e), performed in the presence of hydrogenand of at least one hydrocracking catalyst, at a temperature TR1 ofbetween 250° C. and 480° C., under a pressure of between 2 and 25 MPa,at a space velocity of between 0.1 and 6 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is between 80 and 2000 Nl/l,

Preferably, the hydrocracking step f) according to the invention takesplace at a temperature TR1 of between 320° C. and 450° C., verypreferably between 330° C. and 435° C., under a pressure of between 3and 20 MPa, and very preferably between 9 and 20 MPa, at a spacevelocity of between 0.2 and 3 h⁻¹ and with an amount of hydrogenintroduced such that the litre of hydrogen/litre of hydrocarbon volumeratio is between 200 and 2000 Nl/l.

Preferably, the content of nitrogen in step f), whether this is organicnitrogen dissolved in said unconverted heavy liquid fraction or the NH₃present in the gas phase, is low, preferably less than 200 ppm byweight, preferably less than 100 ppm by weight, more preferably lessthan 50 ppm by weight.

Preferably, the partial pressure of H₂S of step f) is low, preferablythe content of equivalent sulfur is less than 800 ppm by weight,preferably between 10 and 500 ppm by weight, more preferably between 20and 400 ppm by weight.

These operating conditions used in step f) of the process according tothe invention make it possible to maximize the production of “heavynaphtha”; they generally make it possible to obtain conversions perpass, into products having at least 80% by volume of products havingboiling points below 250° C., preferably below 220° C., preferably below190° C. and more preferably below 175° C., of greater than 15% by weightand even more preferably of between 20% and 95% by weight.

In accordance with the invention, the hydrocracking step f) is performedin the presence of at least one hydrocracking catalyst. Preferably, thehydrocracking catalyst of the second step is chosen from conventionalhydrocracking catalysts known to those skilled in the art, such as thecatalysts described above in the hydrocracking step b). Thehydrocracking catalyst used in said step f) may be identical to ordifferent from the one used in step b) and is preferably different.

In one variant, the hydrocracking catalyst used in step f) comprises ahydrogenating-dehydrogenating function comprising at least one noblemetal from group VIII chosen from palladium and platinum, alone or as amixture. The content of metal from group VIII is advantageously between0.01% and 5% by weight and preferably between 0.05% and 3% by weight,the percentages being expressed as weight percentage of oxides relativeto the total mass of catalyst.

Step g)

In accordance with the invention, the process comprises a step g) a stepof hydrogenating at least one portion of the effluent resulting fromstep f) performed in the presence of hydrogen and of a hydrogenationcatalyst, at a temperature TR2 between 150° C. and 470° C., under apressure of between 2 and 25 MPa, at a space velocity of between 0.1 and50 h⁻¹ and with an amount of hydrogen introduced such that the litre ofhydrogen/litre of hydrocarbon volume ratio is between 100 and 4000 Nl/l,said hydrogenation catalyst comprising, and preferably consisting of, atleast one metal from group VIII of the Periodic Table of the Elementschosen from nickel, cobalt, iron, palladium, platinum, rhodium,ruthenium, osmium and iridium alone or as a mixture and not comprisingany metal from group VIB and a support chosen from refractory oxidesupports, and in which the temperature TR2 of the hydrogenation step g)is at least 10° C. below the temperature TR1 of the hydrocracking stepf),

Preferably, said hydrogenation step g) takes place at a temperature TR2of between 150° C. and 380° C., preferably between 180° C. and 320° C.,under a pressure of between 3 and 20 MPa, very preferably between 9 and20 MPa, at a space velocity of between 0.2 and 10 h⁻¹ and with an amountof hydrogen introduced such that the litre of hydrogen/litre ofhydrocarbon volume ratio is between 200 and 3000 Nl/l.

Preferably the litre of hydrogen/litre of hydrocarbon volume ratio ofstep g) is greater than that of the hydrocracking step f).

Preferably, step g) is performed at a temperature TR2 at least 20° C.lower than the temperature TR1, preferably at least 50° C. lower andpreferably at least 70° C. lower.

It is important to note that the temperatures TR1 and TR2 are chosenfrom the ranges mentioned above so as to comply with the deltatemperature according to the present invention, namely that TR2 must beat least 10° C. lower than the temperature TR1, preferably at least 20°C. lower, preferably at least 50° C. lower and more preferably at least70° C. lower.

The technological implementation of the hydrogenation step g) isperformed according to any implementation known to person skilled in theart, for example by injection, in upflow or downflow, of the hydrocarbonfeedstock resulting from step f) and of hydrogen into at least one fixedbed reactor. Said reactor may be of isothermal type or of adiabatictype. An adiabatic reactor is preferred. The hydrocarbon feedstock mayadvantageously be diluted by one or more reinjection(s) of the effluent,resulting from said reactor in which the hydrogenation reaction takesplace, at various points on the reactor, located between the inlet andthe outlet of the reactor, in order to limit the temperature gradient inthe reactor. The stream of hydrogen may be introduced at the same timeas the feedstock to be hydrogenated and/or at one or more differentpoints on the reactor.

Preferably, the metal from group VIII used in the hydrogenation catalystis chosen from nickel, palladium and platinum, alone or as a mixture,preferably nickel and platinum, alone or as a mixture. Preferably, saidhydrogenation catalyst does not comprise molybdenum or tungsten.

Preferably, when the metal from group VIII is a non-noble metal,preferably nickel, the content of metallic element from group VIII insaid catalyst is advantageously between 5% and 65% by weight, morepreferentially between 8% and 55% by weight, and even morepreferentially between 12% and 40% by weight, and even more preferablybetween 15% and 30% by weight, the percentages being expressed as weightpercentage of metallic element relative to the total weight of thecatalyst. Preferably, when the metal from group VIII is a noble metal,preferably palladium and platinum, the content of metallic element fromgroup VIII is advantageously between 0.01% and 5% by weight, morepreferentially between 0.05% and 3% by weight, and more preferably stillbetween 0.08% and 1.5% by weight, the percentages being expressed asweight percentage of metallic element relative to the total weight ofthe catalyst.

Said hydrogenation catalyst may further comprise an additional metalchosen from the metals from group VIII, the metals from group IB and/ortin. Preferably, the additional metal from group VIII is chosen fromplatinum, ruthenium and rhodium, and also palladium (in the case of anickel-based catalyst) and nickel or palladium (in the case of aplatinum-based catalyst). Advantageously, the additional metal fromgroup IB is chosen from copper, gold and silver. Said additionalmetal(s) of group VIII and/or of group IB are preferentially present ina content representing from 0.01% to 20% by weight relative to theweight of the catalyst, preferably from 0.05% to 10% by weight relativeto the weight of the catalyst and even more preferably from 0.05% to 5%by weight relative to the weight of said catalyst. The tin ispreferentially present in a content representing from 0.02% to 15% byweight relative to the weight of the catalyst, so that the Sn/metal(s)from group VIII ratio is between 0.01 and 0.2, preferably between 0.025and 0.055, and even more preferably between 0.03 and 0.05.

The support for said hydrogenation catalyst is advantageously formedfrom at least one refractory oxide preferentially chosen from the oxidesof metals from groups IIA, IIIB, IVB, IIIA and IVA according to the CASnotation of the Periodic Table of the Elements. Preferably, said supportis formed from at least one simple oxide chosen from alumina (Al₂O₃),silica (SiO₂), titanium oxide (TiO₂), ceria (CeO₂), zirconia (ZrO₂) andP₂O₅. Preferably, said support is chosen from aluminas, silicas andsilicas-aluminas, alone or as a mixture. Very preferably, said supportis an alumina or a silica-alumina, alone or as a mixture, and even morepreferably an alumina. Preferably, the silica-alumina contains more than50% by weight of alumina, preferably more than 60% by weight of alumina.The alumina may be present in all possible crystallographic forms:alpha, delta, theta, chi, rho, eta, kappa, gamma, etc., taken alone oras a mixture. Preferably, the support is chosen from delta, theta andgamma alumina.

The catalyst for the hydrogenation step g) may optionally comprise azeolite chosen from Y zeolites, preferably USY zeolites, alone or incombination with other zeolites from beta, ZSM-12, IZM-2, ZSM-22,ZSM-23, SAPO-11, ZSM-48 or ZBM-30 zeolites, alone or as a mixture.Preferably, the zeolite is USY zeolite alone.

Preferably, the catalyst for step g) does not contain zeolite.

A preferred catalyst is a catalyst comprising, and preferably consistingof, nickel and alumina.

Another preferred catalyst is a catalyst comprising, and preferablyconsisting of, platinum and alumina.

Preferably, the hydrogenation catalyst of step g) is different from thatused in the hydrotreating step a) and from those used in thehydrocracking steps b) and f).

The hydrocracking step f) and the hydrogenation step g) mayadvantageously be performed in the same reactor or in differentreactors. When they are performed in the same reactor, the reactorcomprises several catalytic beds, the first catalytic beds comprisingthe hydrocracking catalyst(s) and the following (i.e. downstream)catalytic beds comprising the hydrogenation catalyst(s). In a preferredembodiment of the invention, step f) and step g) are performed in thesame reactor.

The temperature difference between the two steps f) and g) mayadvantageously be managed by one or more heat exchangers or by one ormore quenches (for instance hydrogen or liquid injection quenches) so asto have a temperature with at least 10° C. of difference with thetemperature of step f).

The main objective of the hydrogenation step g) using a hydrogenationcatalyst under operating conditions that are favourable to thehydrogenation reactions is to hydrogenate a portion of the aromatic orpolyaromatic compounds contained in the effluent from step f), and inparticular to reduce the content of HPNA compounds. However, reactionsof desulfurization, of nitrogen removal, of hydrogenation of olefins orof mild hydrocracking are not excluded. The conversion of the aromaticor polyamide compounds is generally greater than 20%, preferably greaterthan 40%, more preferably greater than 80%, and particularly preferablygreater than 90% of the aromatic or polyaromatic compounds contained inthe effluent from step f). The conversion is calculated by dividing thedifference between the amounts of aromatic or polyaromatic compounds inthe hydrocarbon feedstock and in the product by the amounts of aromaticor polyaromatic compounds in the hydrocarbon feedstock (hydrocarbonfeedstock being the effluent from step f) and the product being theeffluent from step g).

In the presence of the hydrogenation step g) according to the invention,the hydrocracking process has a lengthened cycle time and/or an improvedyield of “heavy naphtha”.

Step h)

In accordance with the invention, the process comprises a step h) ofhigh-pressure separation of the effluent resulting from thehydrogenation step g) to produce at least a gaseous effluent and aliquid hydrocarbon effluent.

Said separation step h) advantageously comprises a separation means, forinstance a series of disengagers at high pressure operating between 2and 25 MPa, the purpose of which is to produce a stream of hydrogenwhich is recycled by means of a compressor into at least one of thesteps a), b), f) and/or g), and a hydrocarbon effluent produced in thehydrogenation step g)

Step h) allows the production of a liquid hydrocarbon effluent which isthen recycled to the distillation step d).

Advantageously, said step h) is performed in the same step as step c) orin a separate step.

Step i)

In accordance with the invention, the process comprises a step i) ofrecycling, into said distillation step d), at least one portion of theliquid hydrocarbon effluent resulting from step h).

LIST OF FIGS.

FIG. 1 illustrates one embodiment of the invention.

The VGO-type feedstock is sent a via pipe (1) to a hydrotreating stepa). The effluent resulting from step a) is sent a via pipe (2) into afirst hydrocracking step b). The effluent resulting from step b) is senta via pipe (3) into a high-pressure separation step c) to produce atleast a gaseous effluent (not shown in the FIGURE) and a liquidhydrocarbon effluent which is sent a via pipe (4) into the distillationstep d). The following are drawn off in the distillation step d):

-   -   a gaseous fraction (5),    -   optionally a light petroleum fraction (6) having at least 80% by        volume of products having a boiling point between 20° C. and 80°        C.,    -   a fraction comprising the converted hydrocarbon products having        at least 80% by volume of products boiling at a temperature        below 250° C., (7) and    -   an unconverted liquid fraction having at least 80% by volume of        products having a boiling point above 175° C. (8).

At least one portion of the unconverted liquid fraction containing HPNAsis purged in a step e) via pipe (9).

The purged unconverted liquid fraction is sent via pipe (10) into thesecond hydrocracking step f). The effluent resulting from step f) issent a via pipe (11) into a hydrogenation step g). The hydrogenatedeffluent resulting from step g) is sent a via pipe (12) into ahigh-pressure separation step h) to produce at least a gaseous effluent(not shown in the FIGURE) and a liquid hydrocarbon effluent which isrecycled via pipe (13) into the distillation step d).

EXAMPLES

The examples that follow illustrate the invention without limiting thescope thereof.

Example 1 not in Accordance with the Invention: Basic Case of a Two-StepHydrocracking Process not Comprising a Hydrogenation Step

A hydrocracking unit treats a vacuum gas oil (VGO) feedstock describedin Table 1:

TABLE 1 Type VGO Flow rate t/h 37 Density — 0.92 Initial boiling point(IBP) ° C. 304 Final boiling point (FBP) ° C. 554 S content wt % 2.58 Ncontent ppm by 1461 weight

The VGO feedstock is injected into a preheating step and then into ahydrotreating reactor under the following conditions set out in Table 2:

TABLE 2 Reactor R1 Temperature ° C. 385 Total pressure MPa 14 Catalyst —NiMo on alumina HSV h⁻¹ 1.67

The effluent from this reactor is subsequently injected into a second“hydrocracking” reactor R2 operating under the conditions of Table 3:

TABLE 3 Reactor R2 Temperature ° C. 390 Total pressure MPa 14 Catalyst —Metal/zeolite HSV h⁻¹ 3

R1 and R2 constitute the first hydrocracking step, the effluent from R2is then sent into a separation step composed of a chain for recovery ofheat and then for high-pressure separation including a recyclecompressor and making it possible to separate, on the one hand,hydrogen, hydrogen sulfide and ammonia and, on the other hand, theliquid hydrocarbon effluent feeding a stripper and then an atmosphericdistillation column in order to separate streams concentrated in H₂S, a“Light Naphtha” light petroleum cut (of which 97% by volume of thecompounds have a boiling point of between 27° C. and 80° C.), a “HeavyNaphtha” heavy petroleum cut (of which 96% by volume of the compoundshave a boiling point of between 80° C. and 175° C.) and an unconvertedliquid fraction (UCO) (of which 97% by volume of the compounds have aboiling point above 175° C.). A purge corresponding to 2% by mass of theflow rate of the VGO feedstock is taken as distillation bottoms fromsaid unconverted liquid fraction.

Said unconverted liquid fraction is injected into a hydrocrackingreactor R3 constituting the second hydrocracking step. This reactor R3is used under the following conditions set out in Table 4:

TABLE 4 Reactor R3 Temperature (TR1) ° C. 330 Total pressure MPa 14Catalyst — Metal/zeolite HSV h⁻¹ 2

This second hydrocracking step is performed in the presence of 80 ppm ofequivalent sulfur and 4 ppm of equivalent nitrogen, which originate fromthe H₂S and NH₃ present in the hydrogen and from the sulfur and nitrogencompounds still present in said unconverted liquid fraction.

The effluent from R3 resulting from the second hydrocracking step issubsequently injected into the high-pressure separation step downstreamof the first hydrocracking step and then into the distillation step.

Example 2 in Accordance with the Invention

Example 2 is in accordance with the invention insofar as it is atwo-step hydrocracking process maximizing the production of the “heavynaphtha” fraction in which the effluent resulting from the secondhydrocracking step is sent into a hydrogenation step in the presence ofa hydrogenation catalyst comprising Ni and an alumina support and inwhich the temperature TR2 in the hydrogenation step is at least 10° C.below the temperature TR1 in the second hydrocracking step.

The hydrotreating step in R1, first hydrocracking step in R2 and secondhydrocracking step in R3 are performed on the same feedstock and underthe same conditions as in Example 1. A purge corresponding to 2% by massof the flow rate of the VGO feedstock is also taken as distillationbottoms from the unconverted liquid fraction.

A step of hydrogenation of the effluent resulting from R3 is performedin a reactor R4 downstream of R3. The operating conditions for R4 aregiven in Table 5. In this case, TR2 is 60° C. below TR1.

TABLE 5 Reactor R4 Temperature (TR2) ° C. 270 Total pressure MPa 14Catalyst — Ni/Alumina HSV h⁻¹ 2

The catalyst used in the reactor R4 has the following composition: 28 wt% Ni on gamma alumina.

The hydrogenated effluent resulting from R4 is then sent into ahigh-pressure separation step before being recycled into thedistillation step.

Example 3 in Accordance with the Invention

Example 3 is in accordance with the invention insofar as it is atwo-step hydrocracking process maximizing the production of the “heavynaphtha” fraction in which the effluent resulting from the secondhydrocracking step is sent into a hydrogenation step in the presence ofa hydrogenation catalyst comprising Pt and an alumina support and inwhich the temperature TR2 in the hydrogenation step is at least 10° C.below the temperature TR1 in the second hydrocracking step.

The hydrotreating step in R1, first hydrocracking step in R2 and secondhydrocracking step in R3 are performed on the same feedstock and underthe same conditions as in Example 1. A purge corresponding to 2% by massof the flow rate of the VGO feedstock is also taken as distillationbottoms from the unconverted liquid fraction.

A step of hydrogenation of the effluent resulting from R3 is performedin a reactor R4 downstream of R3. The operating conditions for R4 aregiven in Table 6. In this case, TR2 is 55° C. below TR1.

TABLE 6 Reactor R4 Temperature (TR2) ° C. 275 Total pressure MPa 14Catalyst — Pt/Alumina HSV h⁻¹ 2

The catalyst used in the reactor R4 has the following composition: 0.3wt % Pt on gamma alumina.

The hydrogenated effluent resulting from R4 is then sent into ahigh-pressure separation step before being recycled into thedistillation step.

Example 4 in Accordance with the Invention

Example 4 is in accordance with the invention insofar as it is atwo-step hydrocracking process maximizing the production of the “heavynaphtha” fraction in which the effluent resulting from the secondhydrocracking step is sent into a hydrogenation step in the presence ofa hydrogenation catalyst comprising Ni and an alumina support and inwhich the temperature TR2 in the hydrogenation step is at least 10° C.below the temperature TR1 in the second hydrocracking step.

The hydrotreating step in R1, first hydrocracking step in R2 and secondhydrocracking step in R3 are performed on the same feedstock and underthe same conditions as in Example 1. This time, a purge corresponding to1% by mass of the flow rate of the VGO feedstock is taken asdistillation bottoms from the unconverted liquid fraction.

A step of hydrogenation of the effluent resulting from R3 is performedin a reactor R4 downstream of R3. The operating conditions for R4 aregiven in Table 7. In this case, TR2 is 60° C. below TR1.

TABLE 7 Reactor R4 Temperature (TR2) ° C. 270 Total pressure MPa 14Catalyst — Ni/Alumina HSV h⁻¹ 2

The catalyst used in the reactor R4 has the following composition: 28 wt% Ni on gamma alumina.

The hydrogenated effluent resulting from R4 is then sent into ahigh-pressure separation step before being recycled into thedistillation step.

Example 5: Process Performance

Table 8 summarizes the performance of the processes described inExamples 1 to 4 in terms of “Heavy Naphtha” yield, process cycle timeand overall conversion of the process. The conversion of coronene (HPNAcontaining 7 aromatic rings) performed in the hydrogenation step is alsoreported.

TABLE 8 1 (not in 2 (in 3 (in 4 (in accordance accordance accordanceaccordance with the with the with the with the Examples invention)invention) invention) invention) Scheme R3 alone R3 + R4 R3 + R4 R3 + R4Catalyst in R3 — 28% Ni/ 0.3% Pt/ 28% Ni/ alumina alumina alumina Purge(%) 2 2 2 1 TR1 (° C.) 330 330 330 330 TR2 (° C.) — 270 275 270 Coronene0 95 82 95 conversion (%) (1) “Heavy Naphtha” Base Base Base Base + 1yield point Cycle time Base Base + 7 Base + 4 Base + 5 months monthsmonths Overall 98 98 98 99 conversion (%)

The coronene conversion is calculated by dividing the difference in theamounts of coronene measured upstream and downstream of thehydrogenation reactor by the amount of coronene measured upstream ofthis same reactor. The amount of coronene is measured by high-pressureliquid chromatography coupled to a UV detector (HPLC-UV), at awavelength of 302 nm for which coronene has maximum absorption.

These examples illustrate the advantage of the process according to theinvention which makes it possible to obtain improved performance interms of cycle time, “Heavy Naphtha” yield or overall conversion of theprocess.

The invention claimed is:
 1. A process for producing naphtha from ahydrocarbon feedstock containing at least 20% by volume of compoundsboiling above 340° C., said process comprising the following steps: a) astep of hydrotreating said feedstock in the presence of hydrogen and atleast one hydrotreating catalyst, at a temperature of 200° C. to 450°C., under a pressure of 2 and 25 MPa, at a space velocity of 0.1 to 6h⁻¹ and with an amount of hydrogen introduced such that the litre ofhydrogen/litre of hydrocarbon volume ratio is 100 to 2000 Nl/l, b) astep of hydrocracking at least one portion of the effluent resultingfrom step a), the hydrocracking step b) taking place, in the presence ofhydrogen and at least one hydrocracking catalyst, at a temperature of250° C. to 480° C., under a pressure of 2 to 25 MPa, at a space velocityof 0.1 to 6 h⁻¹ and with an amount of hydrogen introduced such that thelitre of hydrogen/litre of hydrocarbon volume ratio is 80 to 2000 Nl/l,c) a step of high-pressure separation of the effluent resulting fromhydrocracking step b) to produce at least a gaseous effluent and aliquid hydrocarbon effluent, d) a step of distilling at least oneportion of the liquid hydrocarbon effluent resulting from step c)performed in at least one distillation column, from which step thefollowing are drawn off: a gaseous fraction, at least one fractioncomprising the converted hydrocarbon products having at least 80% byvolume of products boiling at a temperature below 250° C., and anunconverted liquid fraction having at least 80% by volume of productshaving a boiling point above 175° C., e) optionally a purging of atleast one portion of said unconverted liquid fraction containing HPNAs,having at least 80% by volume of products having a boiling point above175° C., f) a second step of hydrocracking at least one portion of theunconverted liquid fraction having at least 80% by volume of productswith a boiling point above 175° C., resulting from step d) and/oroptional step e), said step f) being performed in the presence ofhydrogen and of at least one second hydrocracking catalyst, at atemperature TR1 of 250° C. to 480° C., under a pressure of 2 to 25 MPa,at a space velocity of between 0.1 and 6 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is 80 to 2000 Nl/l, g) a step of hydrogenating at least oneportion of the effluent resulting from step f) performed in the presenceof hydrogen and of a hydrogenation catalyst, at a temperature TR2 of150° C. to 470° C., under a pressure of 2 to 25 MPa, at a space velocityof 0.1 to 50 h⁻¹ and with an amount of hydrogen introduced such that thelitre of hydrogen/litre of hydrocarbon volume ratio is 100 to 4000 Nl/l,said hydrogenation catalyst comprising at least one metal from groupVIII chosen from nickel, cobalt, iron, palladium, platinum, rhodium,ruthenium, osmium and iridium alone or as a mixture and not containingany metal from group VIB and a support chosen from refractory oxidesupports, and in which the temperature TR2 is at least 10° C. below thetemperature TR1, h) a step of high-pressure separation of the effluentresulting from the hydrogenation step g) to produce at least a gaseouseffluent and a liquid hydrocarbon effluent, and i) recycling, into saiddistillation step d), at least one portion of the liquid hydrocarboneffluent resulting from step h).
 2. A process according to claim 1, inwhich said hydrocarbon feedstocks are selected from the group consistingof VGOs, vacuum distillates VDs, gas oils, feedstocks originating fromunits for the extraction of aromatics from lubricating oil bases orresulting from the solvent dewaxing of lubricating oil bases,distillates originating from the desulfurization or hydroconversion ofATRs (atmospheric residues) and/or VRs (vacuum residues), or fromdeasphalted oils, feedstocks resulting from biomass, and mixtures of theabovementioned feedstocks.
 3. A process according to claim 1, in whichthe hydrotreating step a) is performed at a temperature of 300° C. to430° C., under a pressure of 5 to 20 MPa, at a space velocity of 0.2 to5 h⁻¹ and with an amount of hydrogen introduced such that the litre ofhydrogen/litre of hydrocarbon volume ratio is 300 to 1500 Nl/l.
 4. Aprocess according to claim 1, in which the hydrocracking step b) isperformed at a temperature of 330° C. to 435° C., under a pressure of 3to 20 MPa, at a space velocity of 0.2 to 4 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is 200 to 2000 Nl/l.
 5. A process according to claim 1, inwhich the following are drawn off from the distillation step d): atleast one fraction comprising the converted hydrocarbon products havingat least 80% by volume of products boiling at a temperature below 190°C., and an unconverted liquid fraction having at least 80% by volume ofproducts having a boiling point above 190° C.
 6. A process according toclaim 1, in which the following are drawn off from the distillation stepd): at least one fraction comprising the converted hydrocarbon productshaving at least 80% by volume of products boiling at a temperature below175° C., and an unconverted liquid fraction having at least 80% byvolume of products having a boiling point above 175° C.
 7. A processaccording to claim 1, in which the hydrocracking step f) is performed ata temperature TR1 of 320° C. to 450° C., under a pressure of 9 to 20MPa, at a space velocity of 0.2 to 3 h⁻¹ and with an amount of hydrogenintroduced such that the litre of hydrogen/litre of hydrocarbon volumeratio is 200 to 2000 Nl/l.
 8. A process according to claim 1, in whichsaid hydrogenation step g) is performed at a temperature TR2 of 180° C.to 320° C., under a pressure of 9 to 20 MPa, at a space velocity of 0.2to 10 h⁻¹ and with an amount of hydrogen introduced such that the litreof hydrogen/litre of hydrocarbon volume ratio is 200 to 3000 Nl/l.
 9. Aprocess according to claim 1, in which step g) is performed at atemperature TR2 at least 20° C. lower than the temperature TR1.
 10. Aprocess according to claim 9, in which step g) is performed at atemperature TR2 at least 50° C. lower than the temperature TR1.
 11. Aprocess according to claim 10, in which step g) is performed at atemperature TR2 at least 70° C. lower than the temperature TR1.
 12. Aprocess according to claim 1, in which the hydrogenation step g) isperformed in the presence of a catalyst comprising nickel and alumina.13. A process according to claim 1, in which the hydrogenation step g)is performed in the presence of a catalyst comprising platinum andalumina.
 14. A process according to claim 1, which is for producingnaphtha from a hydrocarbon feedstock containing at least 80% by volumeof compounds boiling above 340° C.
 15. A process according to claim 1,wherein in step d) the following are drawn off: a gaseous fraction, atleast one fraction comprising the converted hydrocarbon products havingat least 80% by volume of products boiling at a temperature below 175°C., and an unconverted liquid fraction having at least 80% by volume ofproducts having a boiling point above 250° C.
 16. A process according toclaim 1, in which the hydrocracking step f) is performed at atemperature TR1 of 330° C. to 435° C., under a pressure of 9 to 20 MPa,at a space velocity of 0.2 to 3 h⁻¹ and with an amount of hydrogenintroduced such that the litre of hydrogen/litre of hydrocarbon volumeratio is 200 to 2000 Nl/l.
 17. A process according to claim 1, in whichthe hydrogenation step g) is performed in the presence of a catalystconsisting of nickel and alumina.
 18. A process according to claim 1, inwhich the hydrogenation step g) is performed in the presence of acatalyst consisting of platinum and alumina.
 19. A process for producingnaphtha from a hydrocarbon feedstock containing at least 20% by volumeof compounds boiling above 340° C., said process consisting of thefollowing steps: a) a step of hydrotreating said feedstock in thepresence of hydrogen and at least one hydrotreating catalyst, at atemperature of 200° C. to 450° C., under a pressure of 2 and 25 MPa, ata space velocity of 0.1 to 6 h⁻¹ and with an amount of hydrogenintroduced such that the litre of hydrogen/litre of hydrocarbon volumeratio is 100 to 2000 Nl/l, b) a step of hydrocracking at least oneportion of the effluent resulting from step a), the hydrocracking stepb) taking place, in the presence of hydrogen and at least onehydrocracking catalyst, at a temperature of 250° C. to 480° C., under apressure of 2 to 25 MPa, at a space velocity of 0.1 to 6 h⁻¹ and with anamount of hydrogen introduced such that the litre of hydrogen/litre ofhydrocarbon volume ratio is 80 to 2000 Nl/l, c) a step of high-pressureseparation of the effluent resulting from hydrocracking step b) toproduce at least a gaseous effluent and a liquid hydrocarbon effluent,d) a step of distilling at least one portion of the liquid hydrocarboneffluent resulting from step c) performed in at least one distillationcolumn, from which step the following are drawn off: a gaseous fraction,at least one fraction comprising the converted hydrocarbon productshaving at least 80% by volume of products boiling at a temperature below250° C., and an unconverted liquid fraction having at least 80% byvolume of products having a boiling point above 175° C., e) optionally apurging of at least one portion of said unconverted liquid fractioncontaining HPNAs, having at least 80% by volume of products having aboiling point above 175° C., f) a second step of hydrocracking at leastone portion of the unconverted liquid fraction having at least 80% byvolume of products with a boiling point above 175° C., resulting fromstep d) and/or optional step e), said step f) being performed in thepresence of hydrogen and of at least one second hydrocracking catalyst,at a temperature TR1 of 250° C. to 480° C., under a pressure of 2 to 25MPa, at a space velocity of between 0.1 and 6 h⁻¹ and with an amount ofhydrogen introduced such that the litre of hydrogen/litre of hydrocarbonvolume ratio is 80 to 2000 Nl/l, g) a step of hydrogenating at least oneportion of the effluent resulting from step f) performed in the presenceof hydrogen and of a hydrogenation catalyst, at a temperature TR2 of150° C. to 470° C., under a pressure of 2 to 25 MPa, at a space velocityof 0.1 to 50 h⁻¹ and with an amount of hydrogen introduced such that thelitre of hydrogen/litre of hydrocarbon volume ratio is 100 to 4000 Nl/l,said hydrogenation catalyst comprising at least one metal from groupVIII chosen from nickel, cobalt, iron, palladium, platinum, rhodium,ruthenium, osmium and iridium alone or as a mixture and not containingany metal from group VIB and a support chosen from refractory oxidesupports, and in which the temperature TR2 is at least 10° C. below thetemperature TR1, h) a step of high-pressure separation of the effluentresulting from the hydrogenation step g) to produce at least a gaseouseffluent and a liquid hydrocarbon effluent, and i) recycling, into saiddistillation step d), at least one portion of the liquid hydrocarboneffluent resulting from step h).
 20. A process according to claim 19,which is for producing naphtha from a hydrocarbon feedstock containingat least 80% by volume of compounds boiling above 340° C.